Towards the Catalytic Application of a
Monolithic Stirrer Reactor
proefschrift
ter verkrijging van de graad van doctor aan de Technische Universiteit Delft,
op gezag van de Rector Magnificus prof. dr. ir. J.T. Fokkema, voorzitter van het College voor Promoties,
in het openbaar te verdedigen op vrijdag 15 oktober 2004 om 10:30 door Ingrid HOEK
scheikundig ingenieur geboren te Rotterdam
Dit proefschrift is goedgekeurd door de promotor: Prof. dr. J.A. Moulijn
Samenstelling promotiecommissie:
Reactor Magnificus, voorzitter
Prof. dr. J.A. Moulijn, Technische Universiteit Delft, promotor Prof. dr. ir. C.R. Kleijn, Technische Universiteit Delft
Prof. dr. W.F. Hölderich, Rheinisch-Westfälische Technische Hochschule Aachen Prof. dr. F. Kapteijn, Technische Universiteit Delft
Prof. dr. ir. L. Lefferts, Universteit Twente
Dr. ir. A.I. Stankiewicz, Technische Universiteit Delft Dr. ir. T.A. Nijhuis, Universiteit Utrecht
Dit onderzoek is uitgevoerd met financiele steun van de Stichting Technologische Wetenschappen (STW) en DSM N.V.
drukker: Ponsen & Looijen, Wageningen
Table of Contents
Chapter 1 Introduction 1
Chapter 2 Gas-liquid Mass Transfer Characteristics of the Monolithic 9 Stirrer Reactor
Chapter 3 Liquid-solid Mass Transfer Characteristics of the Monolithic 21 Stirrer Reactor
Chapter 4 Power Consumption of the Monolithic Stirrer Reactor 39
Chapter 5 Kinetics of the Etherification of 1-Octanol Catalysed by Zeolite BEA 53
Chapter 6 Etherification of 1-Octanol in the Monolithic Stirrer Reactor 67
Chapter 7 Preparation, Characterisation, and Testing of Palladium on 77 Silica Coated Monoliths
Chapter 8 Hydrogenation of Functionalised Alkynes in the Monolithic 95 Stirrer Reactor
Chapter 9 Summary & Evaluation 109
Samenvatting 117
List of Publications 123
Dankwoord 125
1
1
Structured catalytic reactors
Heterogeneously catalysed gas-liquid reactions are common in the chemical industry. A wide range of reactor types is available for this type of reactions [1]. In the production of fine chemicals the stirred tank reactor is mostly applied, because of its versatility. In such a batch-wise operated reactor the reaction time, catalyst loading, heating, cooling, and mixing can be easily controlled. There are a few drawbacks with respect to the handling of the slurry catalyst and scale-up of such a process. The most important disadvantages are the separation of the catalyst from the reaction mixture, attrition and agglomeration of the catalyst particles, and safety in case of a runaway reaction.
A way to overcome the drawbacks associated with a slurry reactor is to use a structured reactor. In a structured reactor the heterogeneous catalyst is added to the reactor in the form of a well-defined geometry. The main advantages of this type of catalytic reactors are the easy separation of the catalyst from the reaction mixture and the easy scale-up, because the hydrodynamics of a structured catalyst system can be well controlled. Many examples of structured catalyst reactors can be found in literature. A novel structured reactor that is especially suitable for gas-solid reaction is the polylith reactor [2]. The polylith reactor consists of regular elements of catalyst rods that are arranged under a certain angle. The angle can be changed to optimise the pressure drop, mass/heat transfer, and the mixing. Van Hasselt et al. introduced the three levels of porosity reactor as an attractive alternative for a trickle bed reactor in hydrodesulphurisation [3, 4]. The packing in this reactor type consists of permeable catalyst baskets that are spatially arranged in the reactor, resulting in a low pressure drop while the different phases are contacted periodically. Another attractive replacement for a three-phase reactor is the fibrous structured bubble column [5-7]. This reactor consists of a bubble column that is staged with woven catalytic layers. The reactor was used in the hydrogenation of nitrite and 2-butyne-1,4-diol.
A reactor type that receives increasing attention is the monolithic reactor [8-11]. Monolithic structures consist of small parallel channels. The catalyst can be applied onto the walls of the monolith channels. The form of the channels can be round, square or triangular and the material for the walls can be either a metal or a ceramic. The monolithic structures are a good example of a tuneable well-defined catalyst support. The channel diameter, length, and operation conditions can be balanced to meet the demands of a specific process. In Figure 1 monolithic structures of cordierite are shown with different cell densities, in other words different channel sizes.
Monolithic structures are a convenient way to ascertain good contact of the reactants with a sufficient amount of catalyst while a low pressure drop and high mass transfer are reached. These characteristics led to the application of monolithic catalyst in automotive exhaust gas cleaning. Other
Figure 1. Picture of monolithic structures made of cordierite with different cell densities of 200, 400, and 600 cpsi (cells per square inch) [10].
current gas phase applications of the monolithic catalysts are as a selective reduction catalyst in power stations and for ozone destruction in airplanes.
In research on monoliths the focus in the last few decades has been on the use of monolithic catalysts in industrial liquid and gas-liquid processes. In the industrial production of hydrogen peroxide a monolithic reactor is employed for the selective hydrogenation of anthraquinones to their corresponding hydroquinones [12]. This is one of the only known commercial applications of a monolithic reactor in a process involving a liquid phase. Retrofitting an existing slurry reactor into monolithic reactors is an easy way to make the use of the monolithic reactors more attractive for industrial applications. For a process in which a co-current down flow is necessary a loop reactor, in which the liquid and gas are recycled, is most commonly used. For batch operations the monolithic reactor can be placed on an existing slurry reactor as an add-on unit [13]. In this case, the reactor vessel is used as a storage tank. Revamping of a slurry reactor is even easier when the monolithic catalyst beds are placed inside the reactor vessel, either concentric or annular [14, 15]. By inducing gas a liquid circulation is created. The driving force for flow through the monolith channels can be further enhanced by means of mixing device.
2
Catalytic stirrer reactors
Another alternative for revamping a slurry reactor is to combine the mixing and catalytic function by fixing the catalyst to a stirrer. In this way the separation of the catalyst from the reaction mixture becomes easier. Problems with the attrition and agglomeration of the catalyst particles are also avoided. Carberry [16] introduced the stirred basket reactor as a laboratory reactor for the measurement of kinetics under minimal external mass transfer limitations in the gas phase. Figure 2 shows a schematic representation of the stirred basket reactor. The baskets can be filled with catalyst particles of any dimension and shape. By stirring the baskets at high speeds the kinetics of a given reaction can be measured in the absence of axial or radial non-uniformities, and external diffusion limitations.
Figure 2. Schematic representation of a Carberry stirred basket reactor [17].
The stirred basket type reactor is used for measurements of kinetic parameters and to test commercial catalyst pellets in different heterogeneously catalysed reactions. Examples from the literature include: hydrogenation of butynediol [18], the oxidation of phenol [19-22], the oligomerisation of propylene [23], and the catalase enzyme catalysed decomposition of hydrogen peroxide [24].
Kenney and Sedriks [25] compared the activity of a commercial palladium catalyst in the form of pellets with that of the ground pellets. The pellets were used in a stirred basket reactor and the fine
Kawakami et al. [26] determined the reaction characteristics of a catalyst pellet in the hydrogenation of styrene with a stirred basket reactor. They found that internal diffusion limitations of hydrogen played an important role. The reaction rate of the selective hydrogenation of phenylacetylene also proved to be strongly influenced by internal diffusion limitations [27].
The fact that internal mass transfer limitations are not excluded makes the measurement of reaction kinetics with the stirred basket reactor difficult. Various researchers tried to overcome the internal mass transfer limitations associated with pellets, by using more permeable catalyst supports. For instance, a structured catalyst, like a monolithic structure, coated with a thin easy accessible catalyst. Bennet et al. [28] filled spinning baskets with small monolith pieces and mounted different configurations of monolith blocks on a stirrer shaft. They measured the kinetics of the oxidation of propane in the gas-phase. The reaction rates observed for the randomly packed monolithic pieces in the baskets were lower than for the monolithic pieces mounted on the stirrer shaft. The monolith pieces were mounted on the stirrer shaft with the channels in the direction of the flow. The orientation of the channels proved to be important. Liakepoulos et al. [29] used spinning baskets filled with two monolith pieces to determine the gas-phase kinetics of the oxidation of acetaldehyde. The use of thin slices, around 0.4 cm, and sufficient high stirrer speeds, 1860 rpm, resulted in the absence of external mass transfer limitations.
The use of catalytic stirrer reactors in the chemical industry has not received much attention. To overcome the laborious task of the separation of the small catalyst particles from the reaction mixture, Spee suggested to immobilise the catalyst on glass or steel stirrer blades [30]. However, the coated stirrer blades resulted in a small catalytic surface area, resulting in poor yield in the hydrogenation of functionalised alkynes. Schioppa et al. proposed the use of a spinning basket reactor type for the hydrodechlorination in the industrial cleaning of liquid waste streams [31]. The slurry catalyst used in this process tends to agglomerate, resulting in problems during the filtration of the reaction mixture. However, for the use of rather large pellets of catalyst the performance was poor, due to internal mass transfer limitations. A patent claimed the use of a spinning basket reactor filled with monolith pieces in the polymerisation of organosilicon compounds [32]. It was reported that the performance of the spinning basket filled with a monolithic catalyst was promising in comparison with baskets filled with a pelleted catalyst. Furthermore, the patent to use the monolithic pieces directly mounted on the stirrer without the use of wire gauze baskets.
3
Monolithic stirrer reactor
Using a catalytic stirrer reactor that contains structured catalysts as stirrer blades combines the advantages of a structured reactor and a catalytic stirrer. The monolithic stirrer reactor is a reactor in which monolithic structures are mounted on the stirrer shaft as blades. When the stirrer is turned in the liquid a pressure drop is created over the monolith structures. This pressure drop is the driving force for flow through the channels. A monolithic stirrer is a convenient way to transform a slurry reactor in a structured reactor type. In Figure 3 a schematic representation of the monolithic stirrer reactor is shown. This reactor contains six monolithic structures as stirrer blades. There are of course many different stirrer configurations possible, for example by placing more than two monolith blades in one plane.
The main advantage of the monolithic stirrer reactor is the easy separation of the catalyst. Other catalyst handling aspects are also improved, for example attrition and agglomeration of the catalyst is prevented. Another important benefit of the monolithic stirrer reactor compared with a conventional slurry reactor is that the safety is improved. In case of a runaway or an emergency fast shutting down is possible by stopping the impeller.
There are also some disadvantages of the monolithic stirrer reactor. The experience with the monolithic catalyst preparation and catalytic stirrers is limited. While slurry catalysts have been used on industrial scale for decades. Furthermore, there is a limit to the amount of catalyst that can be added to the monolithic stirrer reactor. The amount of catalyst can be tuned by changing the number, length, and cell density of the monoliths. But, the maximum catalyst loading the monolithic stirrer reactor is approximately 4-wt%, while loadings up to 10-wt% can easily be reached in the slurry reactors. In Table 1 a comparison between the novel monolithic stirrer reactor and the conventional slurry reactor is given.
Table 1. Comparison between the conventional slurry reactor with the novel monolithic stirrer reactor.
slurry reactor monolithic stirrer reactor - catalyst separation + +/- safety + + catalyst utilization +/- +/- energy input +/- + catalyst preparation - + catalyst loading - + experience - commercially available catalyst
other can be applied in existing equipment
The size of the monolithic stirrer reactor is limited, mainly with respect to the handling of the monolithic catalysts. This together with the possibility of using the stirrer in existing stirred tanks make the concept in particularly interesting for the fine chemical industry. It is potentially interesting to use as a multi-purpose reactor in the batch-wise manufacturing of a variety of chemicals. A typical way of producing fine chemicals [33].
Edvinsson et al. [34] showed that the principle of a monolithic stirrer reactor concept works for low viscosity liquids. The open structure of the monolithic stirrer blade enables a high throughput of
on the hydrodynamic aspects of the novel reactor type. Subsequently, the use of the monolithic stirrer reactor in a heterogeneously catalysed gas-liquid reaction was demonstrated. It was used to test immobilised homogeneous catalysts in the hydroformylation of 1-octene [35]. The system showed a very stable activity, but the yields obtained with the monolithic stirrer reactor were much lower than in a slurry reactor. A low surface area of the monolithic catalysts and the operating conditions of the reactor mainly caused this difference. Furthermore, the immobilisation of the catalyst was not performed well [36]. Both increasing the partial hydrogen pressure and changing the one-way stirring in periodically changing the stirring direction resulted in higher activities. Apparently, the mixing and/or mass transfer rates were not sufficiently high. Unfortunately, no further investigation of the hydrodynamic aspects of the reactor vessel was performed.
These preliminary studies formed the basis of the project entitled “Development of Monolithic Stirrer Reactors: Operational Characteristics and Scale-up Methods”. The project was a collaboration between the group of prof. dr. Jacob Moulijn (Reactor & Catalysis Engineering) and the group of prof. dr. ir. Harry van den Akker (Kramers Laboratorium for Physical Technology). The scope of the research was to study the hydrodynamic aspects of the monolithic stirrer reactor and to show the catalytic applicability of this new reactor type.
4
Outline of the thesis
The results of the present thesis can be divided in two parts: (i) hydrodynamic characteristics of the monolithic stirrer reactor (Chapter 2, 3, and 4) and (ii) its catalytic feasibility as novel reactor type (Chapter 5, 6, 7, and 8).
The hydrodynamic research is focussed on the mass transfer and power consumption of the monolithic stirrer reactor. The gas-liquid mass transfer characteristics of the monolithic stirrer reactor are discussed in Chapter 2. The different flow regimes are visualised. The effect of the stirrer speed on the volumetric gas-liquid mass transfer coefficient is measured by using the gas absorption method. Chapter 3 deals with the liquid-solid mass transfer characteristics of the monolithic stirrer reactor. The liquid-solid mass transfer coefficient is measured with a fast reaction; the trypsine catalysed hydrolysis of Nα-benzoyl-L-arginine ethyl ester. The influence of the cell density, length of
the monolithic structures, and stirrer speed are also discussed. The channel velocities are calculated with an existing dimensionless model for the liquid-solid mass transfer in capillaries. In Chapter 4 the power consumption of the monolithic stirrer reactor are presented. The influences of the stirrer speed and cell density on the power consumption are investigated.
The second part of the thesis is directed at the applicability of the monolithic stirrer reactor for heterogeneously catalysed liquid and gas-liquid reactions. The etherification of 1-octanol was used as a model reaction for single-phase reactions. The kinetics of this reaction catalysed by zeolite BEA is discussed in Chapter 5. The performance of the monolithic stirrer reactor in the etherification of 1-octanol is proven in Chapter 6. The influences of the cell density and stirrer speed are discussed. The selective hydrogenation of 3-methyl-1-pentyn-3-ol is used to demonstrate the feasiblilty of the monolithic stirrer reactor in two-phase reactions. A palladium catalyst on a silica support catalyses this reaction. First, different preparation methods were studied to incorporate the palladium/silica catalyst on a monolith. Two different coating methods for the application of the silica layer onto the cordierite structures. Chapter 7 describes the influence of the catalyst preparation methods on the performance of the monolithic stirrer reactor in the model reaction. In Chapter 8 the performance of the monolithic stirrer reactor in the hydrogenation of 3-methyl-1-pentyn-3-ol is discussed. The influences of several parameters on the catalytic performance of the novel reactor are shown. The parameters investigated are cell density, stirrer speed, and length of the monoliths. Furthermore, the performance of different stirrer configurations containing 4 monoliths is examined. Finally in Chapter 9 the main conclusions of
5 Literature
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[3] B.W. van Hasselt, D.J. Lindenbergh, H.P. Calis, S.T. Sie, and C.M. van den Bleek, (1997), Chemical Engineering Science, 52 (21-22), 3901-3907
[4] B.W. van Hasselt, P.J.M. Lebens, H.P.A. Callis, F. Kapteijn, S.T. Sie, J.A. Moulijn, and C.M. van den Bleek, (1999), Chemical Engineering Science, 54, 4791-4799
[5] V. Höller, I. Yuranov, L. Kiwi-Minsker, and A. Renken, (2001), Catalysis Today, 69, 175-181 [6] V. Höller, K. Radevik, L. Kiwi-Minsker, and A. Renken, (2001), Industrial and Engineering
Chemistry Research, 40, 1575-1579
[7] E. Joannet, C. Horny, L. Kiwi-Minsker, and A. Renken, (2002), Chemical Engineering Science, 57(16), 3453-3460
[8] R.M. Heck, S. Gulati, and R.J. Farrauto, (2001), Chemical Engineering Journal, 82, 149-156 [9] S. Irandoust and B. Andersson, (1988), Catalysis Reviews – Science and Engineering, 30(3),
341-392
[10] F. Kapteijn, J.J. Heiszwolf, T.A. Nijhuis, and J.A. Moulijn, (1999), CATTECH, 3(1), 24-41 [11] F. Kapteijn, T.A. Nijhuis, J.J. Heiszwolf, and J.A. Moulijn, (2001), Catalysis Today, 66, 133-144 [12] R. Edvinsson Albers, M. Nyström, M. Silverström, A. Sellin, A.-C. Dellve, U. Andersson, W.
Herrmann, and Th. Berglin, (2001), Catalysis Today, 69, 247-252
[13] J.J. Heiszwolf, L.B. Engelvaart, M.G. van den Eijnden, M.T. Kreutzer, F. Kapteijn, and J.A. Moulijn, (2001), Chemical Engineering Science, 56, 805-812
[14] T. Boger, S. Roy, A.K. Heibel, and O. Borchers, (2003), Catalysis Today, 79-80, 441-451 [15] T. Boger, (2002), US Patent 20020081254 assigned to Corning Incorporated, USA [16] J.J. Carberry, (1964), Industrial and Engineering Chemistry, 56(11), 39-46 [17] www.autoclaveengineers.com
[18] F. Turek and H. Winter, (1990), Industrial and Engineering Chemistry Research, 29, 1549-1554 [19] A. Santos, P. Yustos, B. Durban, and F. Garcia-Ochoa, (2001), Industrial and Engineering
Chemistry Research, 40, 2773-2781
[20] A. Santos, P. Yustos, B. Durban, and F. Garcia-Ochoa, (2001), Environmental Science and Technology, 35, 2828-2835
[21] A. Santos, P. Yustos, B. Durban, and F. Garcia-Ochoa, (2001), Catalysis Today, 66, 511-517 [22] A. Santos, E. Barroso, and F. Garcia-Ochoa, (1999), Catalysis Today, 48, 109-117
[23] S. Peratello, M. Molinari, G. Bellusi, and C. Perego, (1999), Catalysis Today, 52, 271-277 [24] T.M. Nguyen and P.F. Greenfield, (1981), Biotechnology and Bioengineering, 23, 805-811 [25] C.N. Kenney and W. Sedriks, (1927), Chemical Engineering Science, 27, 2029-2040
[26] K. Kawakami, S. Ura, and K. Kusunoki, (1976), Journal of Chemical Engineering of Japan, 9(5), 392-396
[27] K. Kawakami and K. Kusunoki, (1976), Journal of Chemical Engineering of Japan, 9(6), 469-474 [28] C.J. Bennet, S.T. Kolaczkowski, and W.J. Thomas, (1991), TransIChemE, 69(B), 209-220 [29] C. Liakepuolos, S. Poulopoulos, and C. Philippopoulos, (2001), Industrial and Engineering
Chemistry Research, 40(6), 1476-1481
[30] M.P.R. Spee, PhD thesis, (1999), University of Utrecht, Utrecht, The Netherlands, 101-116 [31] E. Schioppa, F. Murena, and F. Gioia, (2001), Industrial and Engineering Chemistry Research,
40, 2011-2016
[32] S.T. Kolaczkowski, (1994), European Patent 605 143 A2 assigned to Dow Corning Limited, Great Britain
[33] A. Cybulski, J.A. Moulijn, M.M. Sharma, and R.A. Sheldon, (2001), Fine chemicals manufacture technology and engineering, Elsevier Science B.V., Amsterdam, The Netherlands, chapter 7, 437-520
[34] R.K. Edvinson Albers, M.J.J. Houterman, Th. Vergunst, E. Grolman, and J.A. Moulijn, (1998), AIChE Journal, 44(11), 2459-2464
[35] A.J. Sandee, R.S. Ubale, M. Makkee, J.N.H. Reek, P.C.J. Kamer, J.A. Moulijn, and P.W.N.M. van Leeuwen, (2001), Advances in Synthetic Catalysis, 343(2), 201-206
2
Gas-Liquid Mass Transfer
Characteristics of the
Monolithic Stirrer Reactor
Abstract
The volumetric gas-liquid mass transfer coefficients have been measured for the monolithic stirrer reactor with the gas absorption method with hydrogen in ethanol. Volumetric gas-liquid mass transfer coefficients varying from 0.015 to 0.527 s-1, at stirrer speeds from 200 to 450 rpm, are measured for the standard stirrer configuration containing two monoliths in one plane. The cell density of the monoliths has no effect on the gas-liquid mass transfer. The gas-liquid mass transfer takes place at the interface of the gas and liquid phase. Creating gas bubbles in the liquid increases this interfacial area. In the monolithic stirrer reactor gas bubbles are created by surface aeration. The critical stirrer speed at which gas entrainment takes place is determined visually. It depends on the stirrer and reactor configuration and the physical properties of the gas and liquid. The volumetric gas-liquid mass transfer increases with approximately a factor of three when a stirrer configuration consisting of four monoliths in one plane is used. This is a direct consequence of the decrease in the critical stirrer speed from 230 rpm to approximately 150 rpm. Using four monoliths that are put pair wise on top of each other gave the same results for the volumetric gas-liquid mass transfer coefficient as the standard stirrer configuration.
1 Introduction
Different mass transport processes play a role in the use of the monolithic stirrer reactor in heterogeneously catalysed reactions. The reactants need to be transported to the catalyst layer for reaction to take place. Transport of the reactants takes place via convection, flow of the liquid, and diffusion, molecular transport. Convection of the liquid phase is achieved by stirring. Rotating the monolithic stirrer through the liquid causes a pressure drop over the monoliths. This pressure drop results in convection of the liquid phase, and probably the gas phase, through the monolith channels. On the other hand, at the interface of different phases, molecular transport of matter plays an important role, this process is called mass transfer.
When the monolithic stirrer reactor is used in heterogeneously catalysed gas-liquid reactions it is necessary to dissolve the gaseous reactant in the liquid phase. This gas-liquid mass transfer takes place at the interface of the gas and liquid. The diameter of the reactor determines the interfacial area between the gas and liquid. Gas bubbles in the liquid increase this interfacial area substantially. In the case of the monolithic stirrer reactor discussed in this work gas bubbles are formed by surface aeration. Important parameters that influence the gas-liquid mass transfer coefficient are the physical properties of the liquid and gas phase, and the dimensions of the reactor and stirrer.
The experimental methods for measuring the gas-liquid mass transfer coefficient found in literature can be divided in two groups [1-4]: (i) chemical methods: simultaneous absorption and reaction of gas and (ii) physical methods: absorption or desorption of gas. Literature examples of chemical methods are the absorption of carbon dioxide into alkaline or carbonate buffer solutions [5, 6] and oxygen absorption in a sodium sulphite solution [7]. An advantage of the chemical method is that it allows the determination of the interfacial area as well as the volumetric gas-liquid mass transfer coefficient. There are also a few drawbacks associated with the chemical methods. It is necessary to know the kinetics of the reaction taking place. To ensure that there are no concentration differences the mixing in the reactor needs to be homogeneous. Finally, the values for the gas-liquid mass transfer coefficient obtained with the chemical method are only valid for the same type of liquids. A straightforward method of measuring the volumetric gas-liquid mass transfer coefficient is the gas absorption method [8-12]. With this so-called pressure step method the absorption of a solute gas in a liquid is followed in time. Care must be taken that a sparsely solvable gas is used to minimize changes in the gas and liquid volume during the measurements that influence the volumetric gas-liquid mass transfer measurements.
In this chapter the gas-liquid mass transfer characteristics of the monolithic stirrer reactor are discussed. The flow patterns at different stirrer speeds are described to give insight in the mechanisms taking place in the monolithic stirrer reactor. The volumetric mass transfer coefficient is measured with the gas absorption method. The influence of the stirrer configuration and the stirrer speed on the volumetric gas-liquid mass transfer coefficient of hydrogen in ethanol was investigated.
2 Experimental
2.1 Experimental set-upThe measurements of the volumetric gas-liquid mass transfer coefficient were done in the monolithic stirrer reactor as shown in Figure 1. The reactor had a diameter of 0.20 m while the stirrer diameter measured 0.14 m from tip to tip. The clearance between the bottom of the reactor and the bottom of the monoliths was 0.04 m. The reactor is equipped with a hydrogen and nitrogen supply. Before adding hydrogen the reactor is flushed with nitrogen to assure that no oxygen is present in the reactor system.
sample line cooling water hydrogen vent thermocouple TC nitrogen PC MC 0.12 m 0.20 m 0.14 m 0.01 m
Figure 1. Schematic representation of the monolithic stirrer reactor used for the measurements of the volumetric gas-liquid mass transfer coefficient.
All gas-liquid mass transfer coefficient were measured with ethanol as solvent and hydrogen as the gas phase. The cell densities of the cordierite monoliths used were 50, 200, and 400 cpsi. All monoliths had a diameter of 0.043 m and a length of 0.04 m. The hydrogen pressure during the experiments was 1.7 bara. This pressure was chosen higher than the pressure during reaction, chapter 5 and 6, to increase the gas uptake. A higher gas uptake ensured a larger accuracy of the mass transfer coefficient measurements. The pressure has no influence on the mass transfer coefficient [9, 13]. The pressure controller used was a semiconductor from Kuhlite with a range from 0 to 6 bara. The reactor contains a cooling spiral in the bottom of the reactor to keep the temperature at 298 K during the experiments.
The, so-called, standard stirrer configuration (Figure 1) consisted of 2 monolithic structures. Different stirrer configurations containing 4 monolithic structures, as shown in Figure 2, were tested.
top-view: side-view: top-view: side-view: top-view: side-view:
Figure 2. Schematic representation of the stirrer configurations containing four monolithic structures; a) straight above each other, b) perpendicular above each other, and c) four in one plane.
The configurations containing 4 monolithic structures can be described as follows: 2 monoliths in the standard configuration and 2 monoliths straight above these ones, Figure 2-a, 2 monoliths in the standard configuration and 2 monoliths perpendicular above these ones, Figure 2-b, and 4 monoliths
stirrer configuration. For the standard stirrer configuration and the stirrer containing four monoliths in one plane the reactor was filled with 2.5 kg of ethanol. In the case of the stirrer configurations containing two pair of monoliths above each other the reactor was filled with 4.0 kg of solvent.
2.2 Experimental procedure
The procedure for the measurement of the volumetric gas-liquid mass transfer coefficient consisted of the following steps [8, 9]: (i) degassing at a low pressure until a constant pressure of P0 was reached, (ii) followed by pressurising the reactor to Pm, (iii) when the pressure was constant the
absorption of the hydrogen was followed in time at the desired stirrer speed until a new equilibrium pressure was reached, Pf. Figure 3-a shows a typical results of the course of the pressure against time during the measurement, step three.
1.83 1.84 1.85 1.86 1.87 1.88 0 20 40 60 80 time [s] pressure [bar] 0 1 2 3 4 5 0 20 40 60 time [s] ln[(p m -p f )/(p-p f )] [-]
Figure 3. Typical results of the absorption measurements of hydrogen in ethanol at a temperature of 298 K and a stirrer speed of 300 rpm: a) pressure as a function of the time during the absorption, and b) determination of the gas-liquid mass transfer coefficient.
Integration between the time of the start of the stirrer at Pm and the time at P gives the following
equation for the equilibrium between the gas and liquid phase:
0 0 0
ln
m f m(
)
L f fP
P
P
P
k a t
t
P
P
P
P
−
−
=
−
−
−
[1]In Figure 3-b the ln[(Pm-Pf)/(P-Pf)] is plotted against the time. The volumetric gas-liquid mass
transfer coefficient, kLa, can now be determined from the slope of this plot.
3
Results and Discussion
3.1 Regimes in the monolithic stirrer reactor
The entrainment of gas into the liquid in the monolithic stirrer reactor proceeds via surface aeration. There was no gas-inducing device used, i.e. no gas was directly introduced in the liquid. The surface breakage at which gas entrainment takes place was visualized by using the set-up shown in Figure 4. Bare cordierite monoliths in water were used for this visualisation. The diameter of the reactor was 0.16 m and the stirrer had a diameter of 0.12 m from tip to tip. The stirrer speed was varied from 50 to 300 rpm to show the effect of stirring on the gas entrainment in the liquid.
a) b)
pm
Figure 4. Pictures of the flow patterns in the monolithic stirrer reactor at stirrer speeds varying from 50 to 296 rpm at ambient pressure and room temperature with water and air.
The liquid surface starts moving at a stirrer speed of around 100 rpm. At this point the formation of waves slightly increases the gas-liquid contact area. Above a stirrer speed of 150 rpm a vortex starts to form near the stirrer shaft. At increasing stirrer speed the turbulence of the liquid surface is increased resulting in the formation of gas bubbles at the surface of the liquid, corresponding to a stirrer speed of 200 rpm. The gas bubbles entrained in the liquid at this stage are large and coalesce easily with the surface. Further increase of the stirrer speed results in the break up of this large gas bubbles and makes dispersion of these smaller bubbles in the liquid phase possible. The lowest stirrer speed at which the gas bubbles are dispersed throughout the liquid phase is called the critical stirrer speed, Nc. The critical stirrer speed for the configuration shown in Figure 4 is approximately 220 rpm. At an increase in stirrer speed above the critical stirrer speed the gas bubbles seem to become smaller and more dispersed throughout the liquid.
The critical stirrer speed depends on physical properties of the liquid, like surface tension, density and viscosity, number of baffles, and the configuration of the reactor and stirrer, mainly the distance between the stirrer and gas-liquid interface and the distance between the stirrer and the baffles. The influence of the liquid properties was investigated. Table 1 shows the critical stirrer speeds, determined visually in the different systems.
Table 1. Overview of the critical stirrer speeds for surface aeration for different reactor configurations and physical properties of the liquid.
Conditions Dimensions Critical stirrer speed
solvent gas T [K] p [bara] Hs/Hl [-] Ds/Dr [-] Nc [rpm] water air 298 1.0 0.6 0.75 220 56-wt% water-glycerol air 298 1.0 0.6 0.75 250 ethanol H2 298 1.7 0.6 0.70 230
When the results for the experiments done in water are compared to that of a 56-wt% glycerol water mixture it is clear that the critical stirrer speed depends on the physical properties of the liquid [9]. The experiments done with ethanol and hydrogen are done in the reactor set-up used for the measurements of the volumetric gas-liquid mass transfer. The ratio of the stirrer diameter and reactor diameter, Ds/Dr, is smaller for this reactor system, i.e. the monolithic stirrer blades are closer to the
baffles. This set-up has a critical stirrer speed of 230 rpm when ethanol and hydrogen are used. The influence of this critical stirrer speed on the mass transfer coefficient will be discussed in paragraph 3.2.1.
Another observation made during the visualisation of the different flow regimes is that when the stirrer is stopped gas bubbles can be seen to leave the monolithic structures. This indicates that gas bubbles enter the monolithic channels. The flow pattern of the gas and liquid in the monolith channels is unclear. Different options for the flow regime in the monolithic structures are given in Figure 5 [14-16].
Figure 5. Schematic representation of the possible flow patterns of gas and liquid in the monolith channels; a) dispersed flow, b) Taylor or slug flow, and c) churn flow.
The type of flow depends on a variety of parameters: like the bubble size of the gas in the liquid phase which is related to the stirrer speed, the channel diameter of the monolith, and the physical properties of the liquid and gas phase involved. In the case of the monolithic stirrer reactor it is not expected that surface aeration results in gas bubbles small enough for dispersed flow, Figure 5-a, to take place. The flow regime in the monolithic channels likely resembles the churn flow, Figure 5-c, or the Taylor flow regime, Figure 5-b. For heterogeneously catalysed gas-liquid reactions it was proven that the Taylor flow regime is beneficial for the mass transfer of both the liquid and gas reactants [17-19]. In the Taylor flow regime gas bubbles and liquid slugs flow consecutively through the monolith channels. Because the gas bubble is only separated from the catalyst layer by a thin liquid film direct
gas-solid mass transfer can take place [20]. In the liquid slugs a recirculation pattern is observed. This leads to higher liquid-solid mass transfer in comparison with the laminar flow regime of single-phase flow through the monolith channels.
3.2 Gas-liquid mass transfer coefficient 3.2.1 Influence of the cell density
Figure 6 shows the results of the measurements of the volumetric gas-liquid mass transfer coefficient, kLa, as a function of the stirrer speed for different cell densities. The measurements were done with the standard stirrer configuration containing monoliths with cell densities of 50, 200, and 400 cpsi, corresponding to a channel diameter of 2.98, 1.49, and 1.09 mm, respectively. The average volumetric gas-liquid mass transfer coefficient of at least three measurements is shown together with the 95% confidence interval.
0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 200 250 300 350 400 450 stirrer speed [rpm] kL a [s -1 ]
Figure 6. Average values of the volumetric gas-liquid mass transfer coefficient as a function of the stirrer speed for different cell densities. Experiments performed with the standard stirrer configuration with hydrogen in
ethanol at 298 K and 1.7 bara. Symbols: = 50 cpsi, = 200 cpsi, and = 400 cpsi.
There is no significant difference in the volumetric gas-liquid mass transfer coefficient with varying cell density. This is directly related to the fact that the critical stirrer speed for gas entrainment in the liquid is the same for the different cell densities, namely 230 rpm. When stirrer speeds higher than the critical stirrer speed are used the volumetric gas-liquid mass transfer coefficient increases rapidly. The interfacial area between the gas and liquid phase in the reactor vessel mainly determines the gas-liquid mass transfer coefficient. Gas bubbles entering the monolith channels only influence the gas- liquid-solid and gas-liquid-solid mass transfer characteristics within the channels of the monolith.
The values of the volumetric gas-liquid mass transfer coefficient of the monolithic stirrer reactor are comparable with that of conventional slurry type stirrers used in the laboratory. Dietrich et al. measured the gas-liquid mass transfer coefficient for hydrogen in ethanol at 293 K with a gas inducing Rushton turbine stirrer [8]. The stirrer speed was varied between 750 and 2600 rpm, resulting in a volumetric gas-liquid mass transfer coefficient varying from 0.02 to 1.35 s-1. High stirrer speeds are necessary to obtain gas induction in the liquid. For the gas inducing Rushton turbine a stirrer speed of 1650 rpm gives a mass transfer coefficient of about 0.55 s-1. In the monolithic reactor this gas-liquid mass transfer coefficient is reached at a stirrer speed of 450 rpm.
The gas-liquid mass transfer coefficient is a function of various properties of the gas and liquid phase and the configuration of the reactor and stirrer. Because these parameters were not varied in this experiments no attempts were made to fit a dimensionless correlation for the gas-liquid mass transfer in the monolithic stirrer reactor. The results in this chapter are specific for the hydrogenation reaction discussed in chapter 5 and 6.
3.2.2 Influence of the stirrer configuration
The influence of the stirrer configurations consisting of four monoliths on the volumetric gas-liquid mass transfer coefficient was investigated. First a comparison is made between the stirrer containing four monoliths in one plane, Figure 2-c, and the standard stirrer configuration. Figure 7 shows the average volumetric gas-liquid mass transfer coefficient of three measurements with the corresponding 95 % confidence interval. 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 250 300 350 400 stirrer speed [rpm] kL a [s -1 ]
Figure 7. Influence on the volumetric gas-liquid mass transfer coefficient of the stirrer configuration containing four monoliths in one plane, Figure 2-c, in comparison with the standard configuration. Experiments performed
with hydrogen in ethanol at 298 K and 1.7 bara. Symbols: () = stirrer with four monoliths in one plane, () =
standard stirrer configuration, and the bars represent the 95 % confidence interval.
The error of the measurements with four monoliths in one plane becomes large because the gas absorption takes place in a few seconds. This cannot be accurately measured with the pressure sensor used. The presence of two more monoliths on the stirrer shaft increases the volumetric gas-liquid mass transfer coefficient significantly, approximately a factor of three. This is directly related to the lower stirrer speed for surface aeration. The critical stirrer speed decreases to about 150 rpm when four monoliths are used instead of 230 rpm for the standard stirrer configuration. The two extra monoliths create more turbulence at the gas-liquid surface, facilitating the formation of gas bubbles.
The volumetric gas-liquid mass transfer coefficient was also measured for the stirrer configurations containing four monoliths that are attached to the shaft in pairs on top of each other, Figure 2-a and 2-b. The average volumetric gas-liquid mass transfer coefficient as a function of the stirrer speed is shown in Figure 8 together with the corresponding 95% confidence interval.
0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 250 300 350 400 stirrer speed [rpm] kL a [s -1 ]
Figure 8. Influence on the mass transfer coefficient of the stirrer configurations containing two pairs of monoliths
above each other. Experiments performed with hydrogen in ethanol at 298 K and 1.7 bara. Symbols: () =
perpendicular placement of 2 monoliths above the standard 2 monoliths, Figure 2-a, () = straight placement of
2 monoliths above the standard 2 monoliths, Figiure 2-b, and the bars represent the 95 % confidence interval.
The placement of the 2 monoliths straight or perpendicular above the standard 2 monoliths does not have an influence on the gas-liquid mass transfer in the monolithic stirrer reactor. The values for the volumetric gas-liquid mass transfer coefficient are comparable to the ones found for the standard stirrer configuration. Although the liquid volume is increased to 4.0 kg the placement of the 2 extra monoliths above the standard configuration are capable of sufficient gas entrainment into the liquid. This is in agreement with the values of the critical stirrer speeds for gas aeration, 230 rpm, that is comparable to the value found for the standard stirrer configuration.
4 Conclusions
The volumetric gas-liquid mass transfer characteristics are determined for the monolithic stirrer reactor. Gas bubbles in the liquid enhance the gas-liquid mass transfer substantially. In the monolithic stirrer reactor gas bubbles are created by surface aeration. The critical stirrer speed at which gas entrainment takes place depends on the stirrer and reactor configuration and the physical properties of the gas and liquid phase. The critical stirrer speed for air in water was 220 rpm, while in a mixture of 56-wt% glycerol and water the critical stirrer speed increased to 250 rpm.
The volumetric mass transfer coefficient was measured with hydrogen in ethanol as a function of the stirrer speed. The critical stirrer speed for gas entrainment of this system was 230 rpm. Volumetric gas-liquid mass transfer coefficients varying from 0.015 to 0.527 s-1, at stirrer speeds from 200 to 450 rpm, were found for the standard stirrer configuration containing two monoliths in one plane. The cell density of the monoliths did not have an effect. The volumetric gas-liquid mass transfer was increased with approximately a factor of three when a stirrer configuration consisting of four monoliths in one plane was used. This is a direct consequence of the decrease in the critical stirrer speed from 230 rpm to approximately 150 rpm. Using four monoliths that are put pair-wise on top of each other, both perpendicular and straight, gave the same results for the volumetric gas-liquid coefficient as the standard stirrer configuration.
5 Acknowledgements
Shyreen Dahoe is gratefully acknowledged for her contribution to this work. Corning Inc. is acknowledged for the supply of the monolithic structures.
6 Symbols
Dr = reactor diameter [m]
Ds = stirrer diameter [m]
Hl = height [m] Hs = height [m] kLa = volumetric gas-liquid mass transfer coefficient [s-1]
Nc = critical stirrer speed [rpm] p = pressure [bara] po = degassing pressure [bara]
pf = equilibrium pressure after gas absorption [bara]
pm = equilibrium pressure before gas absorption [bara] T = temperature [K]
t = time [min]
7 Literature
[1] A.W. Patwardhan and J.B. Joshi, (1999), Industrial and Engineering Chemistry Research, 38, 49-80
[2] J.B. Joshi, A.B. Pandit, and M.M. Sharma, (1982), Chemical Engineering Science, 37(6), 813-844
[3] S. Poncin, C. Nguyen, N. Midoux, and J. Breysse, (2002), Chemical Engineering Science, 57, 3299-3306
[4] K. van ‘t Riet, (1979), Industrial and Engineering Chemistry. Process Development and Design, 18(3) 357-364
[5] E. Alper, W.-D. Deckwer, and P.V. Danckwerts, (1980), Chemical Engineering Science, 35, 1263-1268
[6] L.A. Arrua, B.J. McCoy, and J.M. Smith, (1990), AIChE Journal, 36(11), 1768-1772
[7] R. Botton, D. Cosserat, and J.C. Charpentier, (1980), Chemical Engineering Science, 35, 82-89 [8] E. Dietrich, C. Mathieu, H. Delmas, and J. Jenck, (1992), Chemical Engineering Science,
47(13/14) 3597-3604
[9] R.S. Albal, Y.T. Shah, A. Schumpe, and N.L. Carr, (1983), The Chemical Engineering Journal, 27, 61-80
[10] Y.C. Hsu, R.Y. Peng, and C.J. Huang, (1991), Chemical Engineering Science, 52(21/22), 3883-3891
[11] T.I. Mizan, J.Li, B.I. Morsi, M.Y. Chang, E. Maier, and C.P.P. Singh, (1994), Chemical Engineering Science, 49(6), 821-830
[12] V. Linek, P. Benes, and V. Vacek, (1989), Biotechnology and Bioengineering, 33, 1406-1412 [13] M. Teramoto, S. Tai, K. Nishii, and H. Teranishi, (1974), The Chemical Engineering Journal, 8,
223-226
[14] F. Kapteijn, J.J. Heiszwolf, T.A. Nijhuis, and J.A. Moulijn, (1999), CATTECH, 3(1), 24-41
[15] J.J. Heiszwolf, M.T. Kreutzer, M.G. van den Eijnden, F. Kapteijn, and J.A. Moulijn, (2001), Catalysis Today, 69, 51-55
[16] V. Hatziantoniou and B. Andersson, (1982), Industrial and Engineering Chemistry. Fundamentals, 21, 451-456
[18] M.T. Kreutzer, P. Du, J.J. Heiszwolf, F. Kapteijn, and J.A. Moulijn, (2001), Chemical Engineering Science, 56, 6015-6023
[19] M.T. Kreutzer, (2003), PhD thesis, Delft University of Technology, Delft, The Netherlands, 43-67 [20] T.A. Nijhuis, M.T. Kreutzer, A.C.J. Romijn, F. Kapteijn, and J.A. Moulijn, (2001), Chemical
3
Liquid-Solid Mass Transfer
Characteristics of the
Monolithic Stirrer Reactor
Abstract
The liquid-solid mass transfer coefficient in a monolithic stirrer reactor was determined from the rate of the hydrolysis of Nα-benzoyl-L-arginine ethyl ester catalysed by the enzyme trypsine. The trypsine
was immobilised on the monolithic structures by covalent binding. The reaction was verified to be externally mass transfer limited at a temperature above 318 K with the help of an Arrhenius plot.
The influence of the stirrer speed, length of the monolithic structures, and the cell density on the liquid-solid mass transfer coefficient was investigated. Increasing the stirrer speed resulted in an increase in the mass transfer coefficient. When the cell density is reduced, in other words with increasing channel diameter, higher mass transfer coefficients are found. These trends can both be explained by the increasing channel velocity with decreasing cell density and increasing stirrer speed. When the length of the monoliths is reduced it was found that the mass transfer coefficient increased. The developing region for mass transfer plays a role in the short monoliths used.
The channel velocity could be estimated with the help of an existing liquid-solid mass transfer model for monolithic catalysts. The average channel velocity varied between 0.002 and 0.040 m s-1, depending on the cell density of the monolith. With decreasing cell density the average channel velocity increased. For all monolithic structures the channel velocity increased with increasing stirrer speed. The length of the monolithic structures had no effect on the average channel velocity.
1 Introduction
In the monolithic stirrer reactor the reactants need to be transported to the catalyst layer in order for reaction to take place. The catalyst is fixated on the walls of the monolith channels. Liquid-solid mass transfer takes place from the bulk flow in the channels of the monolith to the catalyst layer on the walls of the monolith channels. The pressure drop created over the monolithic structures by using them as stirrer blades causes convection of the reactants through the monolith channels. The channel velocity of the liquid is therefore directly related to the stirrer speed. Other factors that influence the liquid-solid mass transfer are the channel diameter and the length of the monoliths [1].
Different experimental methods for measuring the liquid-solid mass transfer coefficient can be found in literature: (i) dissolving of a solid, like benzoic acid, naphthalene, or cinnamic acid, in a liquid phase [2-6], (ii) electrochemical reactions [7, 8], and (iii) heterogeneously catalysed reactions [9]. The use of soluble solids is a convenient method to measure the liquid-solid mass transfer in the case of packed bed, fluidised bed, or slurry reactors. The benzoic acid is, however, difficult to apply for monolithic structures. With both the dissolving method and the electrochemical method the morphology of the solid surface can change during the measurements, resulting in deviations in the mass transfer coefficient. This problem does not occur when a heterogeneously catalysed reaction is used. In order to measure the liquid-solid mass transfer coefficient it is necessary that the reaction rate is completely externally mass transfer limited. When the reactants reach the catalyst layer they should be converted immediately, so the reactant concentration at the catalyst surface approaches zero. The use of a heterogeneously catalysed gas-liquid reaction makes it sometimes difficult to distinguish between the liquid-solid and gas-liquid mass transfer. By using a liquid phase reaction this problem can be circumvented.
Houterman et al. measured the liquid-solid mass transfer coefficient with the manganese-catalysed decomposition of hydrogen peroxide [10]. Downside of this reaction is that during the reaction oxygen is being formed. The presence of a gas phase might influence the flow pattern in the reactor and monolith channels and, as a consequence, the results might be less reliable.
Horvath and Solomon measured the liquid-solid mass transfer in a tubular reactor with an enzyme reaction: the hydrolysis of Nα-benzoyl-L-arginine ethyl ester to Nα-benzoyl-L-arginine and ethanol, Figure 1 [11]. The reaction was catalysed by trypsine immobilised on the surface of a tube.
HN NH O O CH3 NH2 H NH O + H2O HN NH O NH2 H NH O OH + OH
Figure 1. Hydrolysis of Nα-benzoyl-L-arginine ethyl ester to Nα-benzoyl-L-arginine and ethanol.
The advantage of this reaction is the high activity of the enzyme at moderate conditions. The liquid-solid mass transfer coefficients in this chapter were measured using the trypsine catalysed hydrolysis reaction. The trypsine is immobilised on the monolithic support by covalent binding. The influence of the cell density, i.e. the channel diameter, and the monolith length on the liquid-solid mass transfer coefficient is discussed.
2
Measurements liquid-solid mass transfer
2.1 Immobilisation of the enzyme on the monolithic structure
The enzyme trypsine is immobilised on the monolithic support by a covalent binding. The procedure used for the immobilisation method was developed previously [12]. Anchors need to be created on the surface of the monolith walls before covalent binding of the enzyme can take place. The immobilisation method consists of the following consecutive steps: (i) applying a silica layer to the cordierite structure to increase the number of silanol groups at the surface, (ii) silanisation of the silanol groups with 3-aminopropyl triethoxysilane, (iii) reaction of the amino group of the silane with glutaraldehyde, creating an aldehyde group which forms an anchor for the trypsine, and (iv) covalent binding of an amino group of the enzyme to the aldehyde groups created on the surface of the monolith walls. The trypsine is exclusively immobilised on the surface of the channels inside the monoliths by covering the outside surface with Teflon tape during step ii, iii, and iv.
The silica layer was applied by dip-coating the monoliths in a commercial colloidal silica solution (Ludox AS40, Aldrich, 40-wt% silica in water) for one minute. The monolith channels were emptied with pressurised air. After drying at room temperature the monolithic structures were calcined at 723 K for 4 hours.
For the silanisation reaction, shown in Figure 2, a solution of 15-wt% 3-aminopropyl triethoxysilane (Aldrich) in ethanol was prepared. The pH of a 5-v/v% water-ethanol solution was lowered to 4.5 with acetic acid, after which the 3-aminopropyl-triethoxysilane solution was added. The monoliths were kept in the solution for 4 hours under continuous circulation at room temperature. After rinsing with ethanol the monoliths were dried for 1 hour at 393 K.
Si O OH O Si Si + Si O O Si Si O O NH2 O Si O NH2
Figure 2. Schematic representation of the first step in the immobilisation of trypsine on a silica coated monolith; silanisation of the silanol groups with 3-aminopropyl triethoxysilane.
An anchor for the enzyme can be created by reaction of the amino group of the silane with glutaraldehyde, Figure 3. A 10-wt% solution of glutaraldehyde (Aldrich, 25-wt% solution in water) was prepared in a 0.05 M acetate buffer (pH 5). The buffer was prepared from sodium acetate tri-hydrate (Merck). The monoliths were added to this solution and reaction took place for 2.5 hours at 297 K. After the reaction the monoliths were rinsed with the acetate buffer.
+ O O Si O O Si Si O O NH2 Si O O Si Si O O N O
The binding of the enzyme was realised by reaction between an amino group of the enzyme with the aldehyde group of the formed chain. The reaction was carried out with a solution of 0.2-wt % trypsine in 0.1 M phosphate buffer, Figure 4. The phosphate buffer was prepared from di-sodium hydrogen phosphate di-hydrate (Riedel de Haën). The immobilisation was performed at 278 K while the enzyme solution was continuously circulated through the monoliths.
Si O O Si Si O O N O + H2N Si O O Si Si O O N N
Figure 4. Schematic representation of the third step in the immobilisation of trypsine on a silica coated monolith; covalent binding of the enzyme trypsine to the aldehyde group.
The trypsine concentration during immobilisation could be followed with an UV-spectrophotometer in order to measure the amount of enzyme on the monoliths. After the immobilisation the monoliths were washed with the phosphate buffer and kept in this buffer with 0.1-wt% sodium azide to avoid growth of unwanted micro-organisms at the expense of the enzyme.
2.2 Experimental set-up
The experiments were done in the monolithic stirrer reactor as shown in Figure 5. The reactor had a diameter of 0.16 m while the stirrer diameter measured 0.12 m from tip to tip. The clearance between the bottom of the reactor and the bottom of the monoliths was 0.02 m.
motor baffle monolith blades thermocouple T 0.10 m 0.02 m 0.16 m 0.12 m 0.04 m 0.16 m 0.043 m 0.01 m heating jacket UV- spectro-photometer
Figure 5. Schematic representation of the monolithic stirrer reactor used for the measurements of the liquid-solid mass transfer coefficient.
The reactor was equipped with a heating jacket to keep the reaction mixture at the desired temperature of 319 K. To measure the activation energy of the immobilised enzyme experiments were performed at temperatures of 310, 313, 315, 318, 319, and 320 K with an immobilised ezyme on a 200 cpsi monolith.
All the monolithic structures used in this study were made of cordierite and had a diameter of 0.043 m. The cell densities of the monoliths used are 100, 200, 400, and 600 cpsi (cells per square inch). The characteristics of the different cell densities are given in Table 1.
Table 1. The characteristics of the cordierite monoliths with different cell densities used in the immobilisation of trypsine.
cell density [cpsi] d [mm] a’ [m2 m-3] Aopen [-] 100 2.11 1394 0.69 200 1.49 1945 0.69 400 1.09 2788 0.74 600 0.93 3556 0.80 With increasing cell density the monolith channel diameter decreases, resulting in an increase in total surface area of the monolithic structures. The influence of the length of the monoliths was also investigated. Of the 400 cpsi monoliths the length was varied: 0.01, 0.02, 0.04, and 0.05 m monoliths were used in the measurement of the liquid-solid mass transfer coefficient.
2.3 UV spectrophotometer
A UV spectrophotometer (Thermo Spectronics, UnicamUVseries 5000) was used to determine the concentration of the substrate and product during the reaction. The spectrophotometer was equipped with a tungsten-halogen lamp and a deuterium arc lamp as light sources to cover the wavelength range of 100 to 780 nm. The spectrophotometer is further equipped with a photo multiplier as detector, a flow cell, and an Ebert Monochromator using a 1200 lines/mm Holograph Grating.
A calibration curve was measured to determine the wavelength at which the absorption difference between the substrate and product are maximal. Figure 6 shows the spectra of solutions with different substrate and product concentrations, representing the different levels of conversion between wavelengths of 240 and 290 nm. 0 1 2 3 4 5 240 250 260 270 280 290 wavelength [nm] absorbance [a.u.]
Figure 6. The spectra for different conversion levels showing the absorbance of both the substrate and product in phosphate buffer. The different conversion levels shown are: 0, 25, 50, 75, and 100 %. The total concentration
of substrate and product was always 1 mol m-3. The optimal wavelength was concluded to be 253 nm.
The optimal wavelength for the determination of the concentrations was 253 nm. For the liquid-solid mass transfer coefficient an initial substrate concentration of 0.3 g l-1, corresponding to 0.9 mol m-3, is used for the measurements.
3
Results and Discussion
3.1 Enzyme immobilizationThe enzyme concentration during immobilisation was followed with the UV-spectrophotometer. With a 0.04 m long 600 cpsi monolith the decrease in enzyme concentration was approximately 0.25 g, resulting in 0.50 genzyme gcoating-1. Not all this enzyme is covalently bounded to the surface of the
monolith, so the monoliths were rinsed in the phosphate buffer at 315 K for 2 hours to ensure that no adsorbed enzyme could desorb during reaction. The amount of enzyme that was adsorbed instead of being covalently bonded was small: approximately 0.02 genzyme gcoating-1.
The stability of the immobilised trypsine was studied by performing several experiments spread over 15 days. The immobilised monoliths were stored in the phosphate buffer with 0.1-wt% sodium azide at 278 K in between the measurements. Figure 7 shows the initial activity of the hydrolysis of Nα
-benzoyl-L-arginine ethyl ester for different experiments with the same set of monoliths. The activities have been corrected for the temperature differences between the different experiments.
0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5 5.0 1 2 3 4 5 6 7 8 9 10 test initial activity [g sub m -3 min -1 ]
Figure 7. The initial activity of trypsine immobilised on a 400 cpsi monolith during 10 experiments spread over 15 days. The immobilised monoliths were stored in 0.1-wt% sodium azide in phosphate buffer. Experiments
performed with 0.3 g l-1 substrate in phosphate buffer at a stirrer speed of 100 rpm. The initial activity has been
corrected for temperature differences between the experiments.
The activity of the trypsine immobilised monolithic catalyst showed a slightly decreasing activity upon repetitive use. Experiments 3, 4, and 5 are all done on the 5th day after preparing the immobilised monoliths. During multiple experiments on that day the activity keeps decreasing. This gives an indication that the use of the monoliths causes the deactivation and not the storage in the phosphate buffer with 0.1-wt% sodiumazide. For the measurements of the mass transfer coefficients the catalyst was used for approximately 5 times in a total of 5 days to ensure that deactivation of the enzyme did not play a significant role.
3.2 Mass transfer limited regime
The observed reaction rate of the hydrolysis can be kinetically or externally mass transfer limited. The observed reaction rate depends on the stirrer speed when the reaction rate is externally mass transfer limited. This is not the case when the kinetics determine the observed reaction rate. To determine if the reaction rate is indeed externally limited the data of the reaction rate of hydrolysis at different stirrer speeds is important. Figure 8 shows the evolution of the concentration of substrate during the hydrolysis of Nα-benzoyl-L-arginine ethyl ester catalysed by trypsine immobilised on a
monolith at different stirrer speeds.
0.00 0.05 0.10 0.15 0.20 0.25 0.30 0.35 0 20 40 60 80 100 time [min] conc entra tion [g l -1 ]
Figure 8. Influence of the stirrer speed on the substrate concentration against time. Experiments performed with
a 400 cpsi monolith of 0.05 m in length at an initial substrate concentration of 0.3 g l-1 at 319 K.
At high substrate concentrations the reaction shows a zero order dependency on the substrate concentration and there is not much difference between the stirrer speeds. This indicates that the reaction is mainly kinetically limited in this region. Assuming Michaelis-Menten kinetics the zero order behaviour indicates that the active sites of the enzyme are fully covered. At lower concentrations the reaction shows a first order dependency on the substrate concentration, while the stirrer speed has a large influence. A priori, the first order behaviour can be the result of both kinetical control and external mass transfer control. The former is in agreement with Michaelis-Menten kinetic. However, in the case that the reaction rate depends on the stirrer speed it can be concluded that the reaction is externally mass transfer limited. It was determined that for all the monoliths the reaction showed the strongest dependence on the stirring rate in the substrate concentration region between 0.10 and 0.05 g l-1. For all the calculations the observed reaction rate constant has been determined in this region.
When the reaction rate shows a first order dependency in the substrate concentration, the reaction rate constant can be determined from the slope of the logarithmic normalised concentration against time. Figure 9 shows the first order plots of the experiments shown in Figure 8.
10 rpm 20 rpm
30 rpm 40 rpm
0.1 1.0 0 10 20 30 40 50 60 time [min] C/C 0 [-]
Figure 9. First order plot of the influence of the stirrer speed on the activity of a monolith immobilised with trypsin
in the hydrolysis of Nα-benzoyl-L-arginine ethyl ester; logarithm of concentration from 0.10 to 0.05 g l-1 against
time. Experiments performed with a 400 cpsi monolith of 0.05 m in length at an initial substrate concentration of
0.3 g l-1 at 319 K.
At 10 rpm the reaction rate is clearly externally mass transfer limited. For the stirrer speeds between 20 and 40 rpm the profiles start to coincide, especially at high substrate concentrations. At these high substrate concentrations there is also no true first order behaviour, the lines are not straight in the beginning. Above a stirrer speed of 40 rpm the first order plots of the substrate concentrations become the same. So, external mass transfer limitations are found for stirrer speeds up to 40 rpm in the substrate concentration interval from 0.10 to 0.05 g l-1. Straight lines with differences in the slope are found in the substrate concentration interval used for the other monolithic structures in the range of the stirrer speed as shown in Figures 11 and 12. From the slopes of the lines shown in Figure 9 represent the observed reaction rate constants, kr,obs. The liquid-solid mass transfer coefficient can be
calculated using the following formula:
, L s r obs m
V
k
k
A
=
[1]In which Vl is the liquid volume and Am the geometric surface area of the monolith on which the
trypsine is immobilised. It should be noted that only the geometric surface area of the channel walls is included and not the outer surface area of the monolith pieces.
To verify the conditions for external mass transfer measurements the influence of the temperature on the observed reaction rate constant has been determined. The trypsine is stable up to a temperature of 323 K [10], so the temperature was varied from 310 to 320 K. Figure 10 shows the logarithm of the observed reaction rate constant as a function of the reciprocal of the temperature, measured with a trypsine immobilised 200 cpsi monolith in the hydrolysis of Nα-benzoyl-L-arginine
ethyl ester. From the slope of this Arrhenius plot the activation energy can be determined. The reaction rate constant was determined in the range of 0.10 to 0.05 g.l-1 of the substrate concentration, in which the mass transfer was likely to play a role.
≥ 40 rpm
30 rpm 10 rpm
-4 -3.9 -3.8 -3.7 -3.6 -3.5 3.1 3.12 3.14 3.16 3.18 3.2 3.22 3.24 1000/T [1/K] ln(k obs ) [s -1 ]
Figure 10. Arrhenius plot of a trypsine immobilised 200 cpsi monolith with a length of 0.04 m for the hydrolysis of
Nα-benzoyl-L-arginine ethyl ester. Measurements done at a stirrer speed of 10 rpm with 0.3 g/l substrate in
phosphate buffer solution.
At lower temperatures the activation energy was found to be 40 kJ mol-1, a value also found when free enzymes are used. This is an indication that the reaction is kinetically limited for these conditions. Upon increasing the reaction temperature above 318 K the slope of the Arrhenius plot becomes 5 kJ/mol, confirming that the reaction is externally mass transfer limited. Normally, in the transition from the kinetically towards externally mass transfer limited region the internally limited region is found in the Arrhenius plot. In this region the activation energy should be approximately 20 kJ mol-1, half of the activation energy found in the kinetically limited region. This region is not found for the immobilised trypsine. Several properties of the immobilised system can play a role in this. First of all there is only a very thin silica layer, in which the enzyme could be present. Furthermore, the glutaraldehyde used for the covalent binding of the enzyme functions as a spacer, making the active enzymes flexible. Thus, the most active enzymes are probably the ones that are present on top of, and not inside, the silica layer. The reaction temperature for the liquid-solid mass transfer measurements was set to 319 K.
3.3 Liquid-solid mass transfer coefficient 3.3.1 Influence of the cell density
The influence of the cell density on the liquid-solid mass transfer coefficient is investigated using monolithic structures with a cell density of 100, 200, 400, and 600 cpsi. Figure 11 shows the liquid-solid mass transfer coefficient for the different cell densities at different stirrer speeds.
Ea = 5 kJ/mol
0 1 2 3 4 5 6 7 8 9 0 200 400 600 800
cell density [cpsi]
ks .10 6 [m s -1] 10 20 30 40 50 rpm 0 1 2 3 4 5 6 7 8 9 0 200 400 600 800
cell density [cpsi]
ks .10 6 [m s -1] 10 20 30 40 50 rpm
Figure 11. The liquid-solid mass transfer coefficient for different cell densities as a function of the stirrer speed.
Measurements done in the hydrolysis of Nα-benzoyl-L-arginine ethyl ester at a substrate concentration of 0.3 g l-1
and at 319 K, the length of all monoliths was 0.04 m. Symbols: = 10 rpm, = 20 rpm, = 30 rpm, x = 40
rpm, and + = 50 rpm. Lines are added to guide the eyes.
A general trend for all the cell densities used is that the liquid-solid mass transfer coefficient increases with increasing stirrer speed. Furthermore, it is clear that with increasing cell density, i.e. decreasing channel diameter, the mass transfer coefficient decreases. An obvious interpretation is that the trends are caused by differences in the channel velocity. Kritzinger et al. showed that with increasing stirrer speed and channel diameter the channel velocity of the liquid increases [13].
3.3.2 Influence of the length
To investigate the influence of the length of the monolithic structures the following lengths were used: 0.01, 0.02, 0.04, and 0.05 m. Figure 12 shows the liquid-solid mass transfer coefficient for monoliths with different lengths for different stirrer speeds.
0 1 2 3 4 5 6 7 8 9 10 0 0.01 0.02 0.03 0.04 0.05 0.06 length [m ] ks .10 6 [m s -1] 10 20 30 40 rpm 0 1 2 3 4 5 6 7 8 9 10 0 0.01 0.02 0.03 0.04 0.05 0.06 length [m ] ks .10 6 [m s -1] 10 20 30 40 rpm
Figure 12. The liquid-solid mass transfer coefficient for different lengths of 400 cpsi monolithic blades as a
function of the stirrer speed. Measurements done in the hydrolysis of Nα-benzoyl-L-arginine ethyl ester at a
substrate concentration of 0.3 g l-1 and 319 K. Symbols: = 10 rpm, = 20 rpm, = 30 rpm, and x = 40
